Process for the production of calcium bromide by liquid-liquid extraction

ABSTRACT

A process for the production of calcium bromide from feed brines, particularly from Dead Sea End Brine (EB), is described. The process comprises extracting the feed brine in countercurrent with a composite organic solvent; optionally, purifying the extract to increase the ratio Br:Cl by contacting it with a part of the product; and washing the purified extract with water to yield the product, that is an aqueous solution of CaBr 2 . The composite solvent comprises an anionic extractant, such as an amine or a mixture of amines; a cationic extractant, such as a carboxylic phosphoric or sulphonic acid or a mixtures of said acids; and diluent/modifier, which is an organic solvent. An apparatus for the production of calcium bromide is also described, which comprises: an extraction battery; optionally, a purification battery; and a washing battery, wherein at least one of the batteries comprises a plurality of stages.

FIELD OF THE INVENTION

This invention relates to a process and apparatus for the production ofCaBr₂—calcium bromide—solutions from brines which contain it,particularly from Dead Sea end brine (hereinafter indicated by EB).

BACKGROUND OF THE INVENTION

The term “brine” means herein a concentrated saline solution of sodiumchloride and other salts. Concentrated solution that do not containsodium chloride are not called herein “brines”, although they are socalled at times in the literature. The Dead Sea End Brine hereinafterbriefly indicated by “EB”) is the brine that remains after the finalconcentration stage in the process for the production of carnallite fromDead Sea waters. It contains 20-27 wt % Cl⁻, 0.75-0.95 wt % Br⁻, 2-5 wt% Ca⁺⁺ and 5-7 wt % Mg⁺⁺. Any solution of the above salts, having saltsconcentration substantially within the above limits, is a brinecomprised in this application, regardless of its origin, and what issaid hereinafter about EB should be understood as applying to anysolution having salts concentration substantially within the abovelimits. Calcium bromide solutions, having concentrations e.g. close to50 wt %, are used as clear drilling fluids and optionally may be used asbromine carriers. CaBr₂ is presently prepared by direct reaction ofliquid solution of HBr with Ca(OH)₂ (lime) or CaCO₃, followed byfiltration and concentration. This process, however, is expensive,particularly because hydrobromic acid must be made from bromine, whichis produced from the same EB by stripping with chlorine.

It would be desirable to avoid these chemical processes and to recovercalcium bromide directly from brines which contain it, and this is oneof the purposes of this invention.

Extraction processes are known in the art. Robert R. Grinstead et al.describe the recovery of magnesium chloride from sea water concentratesin the article “Extraction by Phase Separation with Mixed IonicSolvents” in Ind. Eng. Chem. Prod. Res. Develop., 9, No. 1, March 1970.This article describes how magnesium chloride is reversibly extractedfrom an aqueous feed brine, which is a sea water concentrate, by anorganic phase, and is subsequently stripped from the organic phase bycontact with water to produce a magnesium chloride solution. The organicphase used was a solution of a quaternary amine (Aliquat 336) andnaphthenic acid or of a primary amine (Primene JMT) and naphthenic acid,in toluene. Although calcium is present in sea water, little calcium ionis normally found in solution in said concentrates, because of thesubstantial sulfate concentration, and accordingly, it was notconsidered in the said article. The separation considered was mainly ofMgCl₂ from NaCl.

The equipment described by Grinstead et al. consists of an extractorhaving a number of stages and a stripper having another number ofstages. The feed brine is loaded into the extractor, the loaded organicphase is loaded into the stripper. Water is fed to the stripper, fromwhich a product brine is obtained. The organic phase, stripped in thestripper, is returned to the extractor. The final product solution isobtained from the stripper.

C. Hanson et al., in Proceedings International Solvent ExtractionConference 1974, Vol. 1, p. 779-790, have described a process for therecovery of magnesium chloride from sea water concentrates using a mixedionic extractant. Several systems of ionic extractants have been studiedby the authors, and the best results are said to be obtained with anequimolar mixture of Aliquat-336 and Acid-810. Aliquat-336 is a mixtureof quaternary alkyl ammonium chlorides and Acid-810 is essentially amixture of isooctanoic, isononanoic and isodecanoic acids. Toluene wasfound to be a satisfactory diluent.

C. Hanson et al., in “Extraction of Magnesium Chloride from Brines UsingMixed Ionic Extractants”, J. Inorg. Nucl. Chem. 1975, Vol. 37, p.191-198, describe the use of organic extractants, using as aminesAlamine-336 (mixture of tertiary amines), Aliquat-336 (mixture ofquaternary amines), Amberlite LA-2 (mixture of secondary amines) andPrimene JM-T (mixture of isomeric primary amines), and using carboxylicacids as Acid-810, naphthenic acid and Versatic acid 911. The diluentused was toluene.

U.S. Pat. No. 3,649,219 describes a process for extracting inorganicsalts from aqueous solutions which comprises contacting the aqueoussolution with an extractant liquid, which comprises an acid member and abase member, which is an amine, separating the resultinginorganic-salt-containing extractant liquid from the inorganic saltdepleted aqueous phase, and stripping the inorganic salt from theextractant liquid by water.

The prior art deals mainly with separation of MgCl₂ from NaCl. The dataabout the extraction of calcium salts are limited and include mainlychlorides. There are no data on the separation of two salts of divalentmetals. This invention applies to an even more complicated system,dealing with the separation of four salts of divalent metals: CaBr₂,CaCl₂, MgBr₂ and MgCl₂. Furthermore, the references actually describelaboratory experiments and not an industrially valid process, and do nottake into consideration the problems which arise in industrialoperation.

It is therefore a purpose of this invention to provide an industrialprocess for producing calcium bromide by extraction from brines thatcontain it, particularly Dead Sea End Brine.

It is a further purpose of this invention to provide such a processwhich permits to extract calcium bromide in high yields.

It is a still further purpose of this invention to solve the variousproblems which are encountered in the extraction of calcium bromide fromEB and other brines, which will be specifically detailed hereinafter.

Other purposes and advantages of the invention will appear as thedescription proceeds.

SUMMARY OF THE INVENTION

The feed brine considered hereinafter is EB, which contains 20-27 wt %Cl⁻, 0.75-0.95 wt % Br⁻, 2-5 wt % Ca⁺⁺ and 5-7 wt % Mg⁺⁺. However, assaid hereinbefore, the invention applies to other feed brines: ofsimilar composition, no matter what their origin

The process of the invention comprises:

1) extracting a feed brine in countercurrent or crosscurrent with acomposite organic solvent;

2) optionally, purifying the extract to increase the ratio Br:Cl bycontacting it with a part of the product;

3) washing the optionally purified extract with water, preferablydistilled water, to yield the product that is an aqueous solution ofCaBr₂; and

4) optionally, concentrating the product solution to the desiredconcentration, not higher than 65 wt %, e.g. 52 wt %.

The feed of the said process typically has a Br/Cl weight ratio fromabout 1/20 to about 1/40. The product obtained from said processtypically has a Br/Cl weight ratio from about 1:1 to about 1:3, if theoptional purifying step has not been carried out, and from 10/1 to 120/1if it has been carried out.

The aforesaid extraction, purification and washing operations arepreferably carried out in a number of stages.

The composite solvent includes:

an Anionic Extractant: a primary, secondary, tertiary or quaternaryamine with long, straight or branched, aliphatic chains, preferablyC8-C12, or a mixture of said amines;

a Cationic Extractant: a carboxylic acid or alkyl- or aryl-phosphoricacid, or alkyl- or aryl-sulphonic acid, with long aliphatic chain,preferably C8-C12, either straight or branched, or a mixture of saidacids.

a Diluent/Modifier: an aromatic solvent, with short aliphatic chain,preferably C1-C3, e.g., toluene, xylene, anisole, ethylbenzene, etc, ora nitro- or halo-aromatic solvent, such as chlorobenzene, nitrotoluene,or a cycloaliphatic solvent, such as cyclohexane, or an ether or asubstituted aniline.

The composite solvent preferably contains each of the anionic andcationic extractants in molar concentrations in the range of 0.4-0.9 M(moles per liter), the molar concentrations of said anionic and cationicextractants differing from one another by not more than 0.15 M andpreferably being equal.

The composite organic solvent used in the extraction stage is preferablyrefluxed solvent, viz. that resulting from washing the purified extractwith water (hereinafter, “the washed solvent”). It contains smallamounts of bromide, e.g. 200-300 ppm.

The brine resulting from the purification of the extract to increase theratio Br:Cl (hereinafter, “the depleted reflux”) is preferably refluxed,viz. added to the feed brine and fed with it to the extraction stage.

The product solution typically contains 15-35 wt % of CaBr₂.

The brine issuing from the extraction stage. (hereinafter, “the depletedbrine”) is discharged.

The maximum molar loading of the solvent after the extraction by it ofthe feed brine can be equal to the molar concentration of its activecomponents (0.4-0.9 M). The percent loading is the ratio between themoles per liter (M) of salt extracted and the maximum moles per literthat could be extracted (which is equal to the molar concentration ofthe active components of the solvent).

The percent loading of the solvent ranges from about 30% in the winterto 70% in the summer on a molar base. The EB composition changes fromwinter to summer: e.g. the Br wt % may rise from 0.8 to 0.9 and the Clwt % may rise from 23.4 to 26.2. The loading is higher when thetemperature is lower and when the ionic strength is higher. The loadingis a very steep function of temperature and ionic strength.

The apparatus according to the invention comprises: a) an extractionbattery; b) optionally, a purification battery; and c) a washingbattery. All the batteries are preferably multistage. The process can becarried out in various types of equipment. A demonstration plant was runin series of mixer-settlers. In such a plant, each stage comprises amixer in which the organic and aqueous phases are mixed and a settler inwhich they are separated. In later experiments, the process was run in apilot scale pulsed column, where each battery was run in one column. Thecolumn was high enough to yield the requested number of theoreticalstages for each battery. The process can be carried out in other typesof columns e.g. packed columns, Scheibel or Karr columns, just as well.

The yield, calculated as CaBr₂, viz. the ratio between the amount ofcalcium bromide in the product solution to the amount thereof in thefeed brine, is between 20 and 70 wt %. It can be increased by increasingthe number of stages in the various batteries and increasing the ratioof the organic phase to the feed.

BRIEF DESCRIPTION OF THE DRAWINGS

In the drawings:

FIG. 1 is a flowsheet of a process according to an embodiment of theinvention;

FIG. 2 is a schematic block diagram of a battery of mixers and settlers,part of an apparatus for carrying out said process;

FIG. 3 is a schematic diagram of an apparatus comprising a plurality ofbatteries;

FIG. 4 is a schematic illustration of a crosscurrent apparatus.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

The process of the invention is preferably a continuous process, inwhich the organic and aqueous phases are fed countercurrent to oneanother. Thus (see FIG. 2), the upper organic phase from the n-thsettler, is fed to the “n+1” mixer, while the lower aqueous phase fromthe n-th settler is fed to the “n−1” mixer. The organic phase is aclosed circuit all over the plant. Some aqueous phases are introduced orwithdrawn at different stages, as described in the process flowsheet(FIG. 1). The extraction battery comprises, e.g., of 5 mixer-settlers;the purification battery comprises, e.g., 6 mixers-settlers, and thewashing battery also comprises, e.g., 6 mixers-settlers. The height of a40 mm pulsed column for each battery is from 5 to 10 meters.

A pilot plant was erected according to a preferred embodiment of theinvention, which produce, at steady state, 2100 tons/year of CaBr₂ 52 wt%. In said plant, the mixer is a tank of 1250 mm diameter and 1400 mmheight, agitated by a blade turbine. The aqueous and organic flows(streams) enter at the bottom of the mixers by way of separate channels,are pumped to the agitator housing, and are cast forward, generallyperpendicularly.

The settler is a tank with a diameter of 2500 mm and an active height of1800 mm. The flows enter the settler from the mixer by way of arectangular channel to the annulus located in the center of the settler.The annulus is composed of a set of grooves (slits) which creates ahorizontal distribution of the mixture. The separation region in thesettler is composed of 12 units of stream distributors. The calmingregion is built from horizontal leaves through which the mixture passesand is separated. At the upper edge of the settler, there is a channelfor the perimetric collection, with an exit opening of 8 inches to thenext mixer. At the bottom of the settler, there is a directing system byway of a variable leg, which determines the aqueous phase height in thesettler.

The material of the construction of the mixers and the settlers isF.R.P. Cristic 600 PA or similar.

Mixer-Settlers that can be used to carry out the invention are known inthe art and are described for instance in U.S. Pat. No. 3,489,526.Compact settlers that can be used to carry out the invention are knownin the art and are described for instance in U.S. Pat. No. 3,563,389.Turbine mixers that can be used to carry out the invention are known inthe art and are described for instance in U.S. Pat. No. 3,973,759.

At least four different types of columns, that can be used to carry outthe invention, are known in the art: packed columns, described forinstance in G. Stevens in J C Godfrey and M. Slater “Liquid-liquidExtraction Equipment, BPC (Bateman Pulsed Column), Karr reciprocatingcolumn, described for instance in A E Karr and T C Lo, Chem. Eng. Prog.72, 68 (1976) and Scheibel agitated column, described for instance inU.S. Pat. No. 3,389,970.

The process of the invention comprises the extraction of the feed,typically EB, by a composite solvent, which includes an anionicextractant, a cationic extractant and a diluent, which serves also as amodifier.

As set forth hereinbefore, the process may be carried out in anycomposite solvent which contains the following components:

Anionic Extractant: a primary, secondary, tertiary or quaternary aminewith long, straight or branched, aliphatic chains, preferably C8-C12, ora mixture of said amines.

-   -   Cationic Extractant: a carboxylic or alkyl- or aryl-phosphoric,        or alklyl-or aryl-sulphonic acid, with long aliphatic chain,        preferably C8-C12, either straight or branched, or a mixture of        said acids.    -   Diluent/Modifier: an aromatic solvent, with short aliphatic        chain, preferably C1-C3, e.g., toluene, xylene, anisole,        ethylbenzene, etc, or a nitro- or halo-aromatic solvent, such as        chlorobenzene, nitrotoluene, or a cycloaliphatic solvent, such        as cyclohexane, or an ether or a substituted aniline.

The anionic and cationic extractants should have a similar molarconcentration in the solvent mixture, in the range of 0.4-0.9 M,preferably equal molar concentrations or molar concentrations thatdiffer by no more than 0.15 M. A preferred concentration is about 0.8 M.Below 0.4 M, the loading of the solvent dropped below reasonable values,while above 0.9 M, the solvent has a high viscosity that hinders theseparation of the organic phase from the aqueous phase.

The selectivities of Br versus Cl, indicated herein as S(Br,Cl), and ofCa versus Mg, indicated herein as S(Ca,Mg), are defined as:

${S\left( {{Br},{Cl}} \right)} = \frac{\frac{\left\{ {Br} \right\}{org}}{\left\{ {Br} \right\}{aq}}}{\frac{\left\{ {Cl} \right\}{org}}{\left\{ {Cl} \right\}{aq}}}$viz. the ratio of the ratio of the molar or weight concentration of Brin the organic phase to that in the aqueous phase to the ratio of themolar or weight concentration of Cl in the organic phase to that in theaqueous phase;

${{S\left( {{Ca},{Mg}} \right)} = \frac{\frac{\left\{ {Ca} \right\}{org}}{\left\{ {Ca} \right\}{aq}}}{\frac{\left\{ {Mg} \right\}{org}}{\left\{ {Mg} \right\}{aq}}}},$viz. the ratio of the ratio of the molar or weight concentration of Cain the organic phase to that in the aqueous phase to the ratio of themolar or weight concentration of Mg in the organic phase to that in theaqueous phase.

The selectivities are not functions of the concentrations, and theloading is roughly linear to the concentration of the extractants.S(Br,Cl) is about 8 to 30, typically 12-18, and S(Ca,Mg) is so high thatthe extractability of magnesium is negligible for all practicalpurposes.

The solvent used in the following Examples consisted of:

0.8 M. (about 43 wt %) Alamine 336 (mixture of tridecyl and tridecylamines).

0.8 M (about 12 wt %) Cekanoic Acid (isodecanoic acid).

About 45 wt % Xylene (mainly mixture of metaxylene and ethylbenzene)

The solvent that comes into actual contact with the feed is a refluxed,washed solvent, that contains small amounts of bromide, as set forthhereinbefore. The aforesaid data refer to the unloaded, originalsolvent.

FIG. 1 is a flowsheet of a process according to an embodiment of theinvention which comprises the purification stages. Numerals 10, 11 and12 indicate respectively the extraction, purification and washingprocess stages. Line 13 leads the extract from extraction topurification. Line 14 leads the purified extract from purification towashing. Line 15 indicates the reflux of the washed solvent toextraction. Line 16 indicates the water feed. The product solution isdrawn from the washing as indicated at 17, and is divided into a productreflux to purification, as indicated at 18, and a final productsolution, drawn from the process as indicated at 19. Using the productas reflux (stream 18), the purification stage yields a purified extract,that has a ratio Br:Cl from 10:1 to 120:1. Said extract has a highconcentration of Br⁻(above 2 wt %), which leads to high entrainment ofthe aqueous phase in the organic phase after separation in the laststages of the purification battery. Up to 3 vol % entrainment isobtained, while the maximum allowed to avoid back mixing is. 0.1 vol %.

In order to lower the aforesaid concentration of Br⁻, the flowsheet ofFIG. 1 was modified, in a preferred embodiment of the invention, by theaddition of water in an intermediate point of the purification processstage, as indicated at 20 in FIG. 1. It lowered the ionic strength ofboth the aqueous and the organic phases, and sharply decreased theaqueous phase entrainment, without any harm to the yield or theproduction rate.

A problem that can arise is constituted by the leakage of solvent fromthe plant. This leakage is mainly due to the presence of organic solventin the depleted brine, causes an ecological problem, and leads to lossof solvent. The components of the solvent have different solubilities inthe brine and, on the other hand, there is solvent entrainment in thedepleted brine, consisting of drops of solvent that did not separate inthe settler. The solubilities of the organic component in the brine areas follows: alamine —0.2 ppm; cekanoic—20 ppm. Thus the solvent lossescaused by solubility are negligible and cannot be prevented, anyhow.

Factors that contribute to increase the leakage caused by entrainmentare:

a) The decrease in the temperature of the brine and of the solvent,which increases the viscosity and decreases the phase separation rate.This phenomenon is particularly important in winter.

b) An increase in the concentration of surface-active agents whichaccumulate in the solvent also increase the entrainment of solvent inthe depleted brine.

c) Accumulation of deposits and gels in the settlers decrease thecapability of phase separation.

The leakage problem is solved, in an embodiment of the invention, byphysical separation of the solvent entrained in the depleted brine.Physical separation can be effected e.g. by the use of a separator witha filling of particulate material, specifically sand, having agranulometry of 0.8-1 mm. This sand filter, in an embodiment of theinvention, had a diameter of 3.2 m. and the sand bed height was 1 m. Theflow rate through it was 3-4 m/h and every 24 hours it wascounter-currently washed with water at a flow rate of 35-40 m/h for15-20 min.

FIG. 2, which is self-explanatory, schematically shows how the aqueousphase (viz., the brine) flows through successive couples of mixers andsettlers, while the organic phase (viz., the solvent) flows through thesame in countercurrent.

FIG. 3 schematically illustrates how the aqueous and the organic phaseflow successively though several extraction stages, while flowing incountercurrent in each stage.

FIG. 4 schematically illustrates a crosscurrent apparatus, in which theorganic phase flows successively through several stages, and iscontacted in each stage with a fresh stream of aqueous phase.

The following examples describe operations that were carried out in thepilot plant hereinbefore described.

Example 1

The following are the representative results of a summer run (August)according to the flowsheet of FIG. 1. The following parameters obtained:

Brine composition: 0.90% Br—, 26.2% Cl—

Flow rates: 43 m³/hr end brine, 85 m³/hr solvent, 3.6 m³/hr wash water.

Production rate: 1110 lt/hr of 18.5% CaBr₂ (490 kg of 52% CaBr₂, with0.7% Cl— in the 52% product)

Yield of Br— extraction: 37%

Representative concentrations:

Depleted brine: 0.53% Br—, 24.4% Cl—

Extract: 1.2% Br—, 1.2% Cl—

Washed solvent 0.03% Br₂

The percentages, here and hereinafter, are by weight unless otherwiseindicated.

Stream Flow Rate Stream No. m³/h wt % Cl- wt % Br- E.B. 1 43 26.2 0.90W.S. 15 85 0 0.03 Water to washing 16 3.6 0 0 R 18 2.3 0.25 14.2 Waterto purification 20 2.2 0 0 Pr 19 1.11 0.25 14.2 DB 2 — 24.4 0.53 Ex 13 —1.2 1.2 P. Ex 14 — 0.005 0.7 DR 3 — 16.7 8.6

Example 2

Representative results of a winter run (January) from the same plant aregiven below. The yield of CaBr₂ ⁻ extraction was 23 wt %.

Stream Flow Rate Stream No. m³/h wt % Cl- wt % Br- E.B. 1 35 23.4 0.80W.S. 15 50 0 0.03 Water to washing 16 2.0 0 0 R 18 1.8 0.25 14.3 Waterto purification 20 1.2 0 0 Pr 19 0.38 0.25 14.3 DB 2 — 22.1 0.60 Ex 13 —1.4 0.86 P. Ex 14 — 0.005 0.77 DR 3 — 16.8 5.2

Example 3

This Example refers to a summer run (temperature 50° C.) carried out ina pilot that consists of nine stages of mixer settler, 45 1/hr of DeadSea EB met, in countercurrent flow, 85 1/hr of washed solvent. Theextract was washed with 6 1/hr of distilled water, without anypurification. Five stages were used for extraction and four for washing.

The inlet brine contained 26.7 wt % Cl, 0.92 wt % Br and 4.08 wt % Ca⁺⁺and 6.8 wt % Mg⁺⁺. The product had 5.2 wt % Br⁻ and 17.1 wt % Cl⁻. 85 wt% of the Br⁻ in the feed was extracted.

Example 4

Winter conditions—temperature of 25° C.

The plant and the flow rates are as in Example 3, but the composition ofthe feed is 24 wt % Cl⁻, 0.84 wt % Br⁻, 3.25 wt % Ca⁺⁺ and 6.8 wt %Mg⁺⁺.

The composition of the product was 5.2 wt % Br⁻ and 15.5 wt % Cl⁻. 92 wt% of the Br⁻ in the feed was extracted.

While embodiments of the invention have been described for the purposeof illustration, it will be apparent that the invention can be carriedinto practice with many modifications, variations and adaptations,without departing from its spirit or exceeding the scope of the claims.

1. A process for the production of a calcium bromide solution comprisingthe steps of: a) extracting a feed brine comprising Cl⁻, Br⁻, Ca⁺⁺ andMg⁺⁺ countercurrently with a composite organic solvent in a firstplurality of stages, thereby obtaining an extract and a depleted brine;and b) washing said extract with water to yield a product of an aqueoussolution of CaBr₂ and a refluxed solvent, wherein said composite organicsolvent comprises (1) an anionic extractant selected from the groupconsisting of primary, secondary, tertiary, and quaternary amines, andmixtures thereof having long, straight, or branched aliphatic chains (2)a cationic extractant selected from the group consisting of carboxylicacids, alkyl- or aryl-phosphoric acids, and alkyl- or aryl-sulphonicacids, and mixtures thereof having long, straight, or branched aliphaticchains and (3) a diluent-modifier selected from the group consisting ofaromatic solvents having short aliphatic chains, nitro- or halo-aromaticsolvents, and cycloaliphatic solvents.
 2. The process according to claim1, further comprising, prior to step b), a purification step forpurifying said extract in a second plurality of stages to increase theBr:Cl ratio by contacting said extract with a part of said product,thereby obtaining a purified extract and a depleted reflux.
 3. Theprocess according to claim 2, wherein said product is concentrated to aconcentration of about 52% by weight CaBr₂.
 4. The process according toclaim 1, wherein said feed brine has a Br:Cl weight ratio from about1:20 to about 1:40, and said product has a Br:Cl weight ratio from about10:1 to about 120:1.
 5. The process according to claim 1, wherein saidbrine and said composite organic solvent flow in countercurrentlythrough said first plurality of stages.
 6. The process according toclaim 1, wherein said brine flows successively through said firstplurality of stages and said composite organic solvent flowscountercurrently in each of said first plurality of stages.
 7. Theprocess according to claim 1, wherein said composite organic solventflows successively through said first plurality of stages and iscontacted in each of said first plurality of stages with a fresh streamof brine.
 8. The process according to claim 1, wherein thediluent-modifier is selected from the group consisting of cyclohexane,ethers, substituted anilines, toluene, xylene, anisole, ethylbenzene,chlorobenzene and, nitrotoluene.
 9. The process according to claim 1,wherein said composite organic solvent contains each of said anionic andcationic extractants in molar concentrations in the range of 0.4-0.9 M(moles per liter), and the difference in molar concentrations betweensaid anionic and said cationic extractants is not more than about 0.15 M(moles per liter).
 10. The process according to claim 1, wherein saidcomposite organic solvent used to extract said feed brine is saidrefluxed solvent resulting from washing said extract of said step b).11. The process according to claim 2, wherein said depleted reflux isadded to said feed brine and fed with it to said extracting step. 12.The process according to claim 1, wherein said process is continuous.13. The process according to claim 2, further comprising adding water inan intermediate point of said purification step.
 14. The processaccording to claim 1, further comprising concentrating said product to aconcentration of about 65 wt % or less.
 15. The process according toclaim 1, further comprising separating an entrained solvent from saiddepleted brine.
 16. The process according to claim 1, wherein said feedbrine is Dead Sea End Brine (EB).